Two-stage method for producing butanediol with intermediated separation of succinic anhydride

ABSTRACT

Optionally alkyl-substituted 1,4-butanediol is prepared from C 4 -dicarboxylic acids and/or of derivatives thereof by:
         a) a gas stream of the C 4 -dicarboxylic acid or the derivative thereof in a first reactor in the gas phase to obtain a product which contains mainly optionally alkyl-substituted γ-butyro-lactone;   b) removing succinic anhydride from the product of step a);   c) catalytically hydrogenating the product of step b) in a second reactor in the gas phase to obtain optionally alkyl-substituted 1,4-butanediol;   d) removing the desired product from intermediates, by-products and any unconverted reactants; and   e) optionally recycling unconverted intermediates into one or both hydrogenation stages.       

     The catalysts employed in each of the hydrogenation stages comprise ≦95% by weight of CuO, and ≦5% by weight of an oxidic support, and the second reactor has a higher pressure than the first reactor.

The present invention relates to a two-stage process for preparingoptionally alkyl-substituted butanediol by catalytic gas-phasehydrogenation of substrates which are selected from the group consistingof derivatives of maleic acid and succinic acid and also these acidsthemselves. For the purposes of the present invention, derivatives areanhydrides which, like the acids, may have one or more alkylsubstituents. Removal of succinic anhydride after the firsthydrogenation stage allows activity, selectivity and lifetime of thecatalyst in the second hydrogenation stage to be increased.

The hydrogenation of MA which is known per se leads via the intermediatesuccinic anhydride (SA) initially to γ-butyrolactone (GBL). Furtherhydrogenation then leads to tetrahydrofuran (THF), n-butanol (BuOH)and/or n-butane. GBL and butanediol (BDO) are in an equilibrium whichcan be shifted by suitable measures substantially to the side ofbutanediol. However, butanediol can react just as easily as GBL byoverhydrogenation to give butanol and butane; cyclization of butanediolgives THF. These products cannot be converted back to BDO or GBL. WhenBDO is the desired product, the formation of THF in particular has to beavoided.

The gas phase hydrogenation of purified maleic anhydride (MA) tobutyrolactone (GBL) and the conversion of purified GBL to BDO are tworeactions which have been known for many years. To carry out these twocatalytic reactions, the literature describes numerous catalyst systems.Depending on the composition of the catalysts and the reactionparameters chosen, such catalysts give different product distributions.Processes for direct preparation of butanediol starting from MA arelikewise already known.

When GBL and BDO which have alkyl substituents are to be prepared, thereis the possibility of using the corresponding alkyl-substituted speciesof the abovementioned reactants.

The catalysts used for hydrogenating MA to one of the abovementionedproducts, in particular in older processes, frequently contain chromium.This is reflected by the patent literature where a large number ofpatents and patent applications disclose the use of chromium catalystsfor the hydrogenation reaction, although the hydrogenation in most casesis restricted to MA as the reactant.

The documents hereinbelow describe the use of chromium catalysts forhydrogenating MA.

EP-A 0 322 140 discloses a continuous process for preparingtetrahydrofuran (THF) and for coproduction of THF and GBL by gas phasehydrogenation of MA and SA. The claimed catalyst contains copper, zincand aluminum and a further element of groups IIA, IIIA, VA, VIII, IIIBto VIIB, the lanthanide and actinide series, and also Ag and Au. At 40bar, these catalyst systems achieve THF yields of 90-95% starting frompure MA, and at a pressure of about 20 bar mixtures of GBL and THF canbe obtained.

In U.S. Pat. Nos. 4,965,378 and 5,072,009, a similar catalyst is usedwhich may, however, additionally contain Si, Ge, Sn and Pb. The use ofsuch catalysts results in high quantities of THF (from 95% to 31.4%)which cannot be converted to butyrolactone or butanediol.

EP-A 0 404 408 discloses an MA hydrogenation catalyst whosecatalytically active material corresponds substantially to the materialof U.S. Pat. No. 5,072,009. It is used fixed on a support as a coatedcatalyst. In the examples, exclusively chromium catalysts are used. HighGBL yields can be realized at a pressure of 2 bar, but when a higherpressure is used the THF yield increases while the GBL yield decreases.

U.S. Pat. No. 5,149,836 discloses a multistage gas phase process forpreparing GBL and THF with variable product selectivities by, in a firststage, passing a mixture of pure MA and hydrogen over a catalyst whichcomprises copper, zinc and aluminum. This crude reaction effluent isthen passed over a chromium catalyst to prepare THF.

WO 99/38856 discloses a catalyst comprising only copper and chromiumwhich allows GBL selectivities of from 92 to 96 mol % to be obtained instraight paths starting from pure MA.

EP-A 638 565 discloses a copper-, chromium- and silicon-containingcatalyst which has a composition in one example of about 78% of CuO, 20%of Cr₂O₃ and 2% of SiO₂. Using pure MA and nitrogen-hydrogen mixtures,GBL yields of 98% could be obtianed.

The documents hereinbelow disclose the use of chromium-free catalystsfor hydrogenating MA.

GB-A 1 168 220 discloses a gas phase process for preparing GBL byhydrogenating MA or SA over a binary copper-zinc catalyst to GBL. In allexamples, operation is effected at atmospheric pressure and GBL yieldsof 94 mol % could be obtained starting from pure MA.

DE-A 2 404 493 likewise discloses a process for preparing GBL bycatalytically hydrogenating mixtures of MA, SA, maleic acid, succinicacid and water over metallic catalysts, and as well as copper chromitecatalysts, copper-zinc and copper-zinc-aluminum precipitated catalystsare also used.

WO 91/16132 discloses the hydrogenation of MA to GBL using a catalystcomprising CuO, ZnO and Al₂O₃ which is reduced at from 150° C. to 350°C. and activated at 400° C. The activation is intended to increase theon-stream time of the catalyst system.

A catalyst comprising CuO and ZnO is disclosed by U.S. Pat. No.6,297,389. After activation, this converts pure MSA to GBL in yields offrom 92 to 96% and in straight paths starting from pure MA.

WO 95/22539 discloses a process for preparing GBL by catalyticallyhydrogenating MA and/or SA over a catalyst which consists of copper,zinc and zirconium. Starting from pure MA, GBL yields of up to 99% areobtained.

WO 99/35136 discloses a two-stage process for preparing GBL and THF byhydrogenating MA in a first stage using a copper catalyst and passingthis reaction effluent over an acidic silicon-aluminum catalyst.

WO 97/24346 describes a copper oxide-aluminum oxide catalyst whichhydrogenates MA to GBL in yields of 92 mol %.

The conversion of GBL to BDO is also a reaction which has been known forsome time. The documents mentioned hereinbelow disclose this reactionusing chromium catalysts.

DE 1 277 233 discloses a process for preparing mixtures of differentalcohols by hydrogenating lactones using hydrogen. The catalysts usedare copper chromite admixed with barium on an inactive aluminum oxidesupport.

GB-A 1 230 276 discloses a process for preparing BDO from GBL over acopper oxide-chromium oxide catalyst at a temperature of from 180° C. to230° C.

According to DE-A 2 231 986, copper chromite catalysts which are dopedusing potassium, sodium, rubidium, aluminum, titanium, iron, cobalt ornickel increase the on-stream time of the catalysts.

According to DE-A 2 501 499, BDO is prepared using a mixture of dioxane,GBL, water and carboxylic acids. The reaction described takes place athigh pressure (170 bar) in the liquid phase, preferably using thesolvent dioxane, and copper-chromium oxide catalysts are likewise used.

Copper chromite catalysts are doped according to J01121-228-A using Pdin order to achieve a higher conversion.

Further copper chromite catalysts are described by Dasunin, Maeva, Z.Org. chim. 1 (1965), No. 6, p996-1000; JA 5366/69; JA 7240770; J49024-906; J49087-610; and the examples concern the liquid phaseconversion of pure GBL to BDO.

The gas phase hydrogenation of pure GBL to butanediol over copperchromite catalysts is described in U.S. Pat. No. 4,652,685. At apressure of 41 bar and a conversion of 60-68%, a BDO selectivity of92-97% could be achieved.

U.S. Pat. Nos. 5,406,004 and 5,395,990 disclose processes for preparingmixtures of alcohols and diols by hydrogenating pure GBL over coppercatalysts. A hydrogenation zone filled with copper catalyst is chargedwith hydrogenation feed and hydrogen at a temperature of from 150 to350° C. and pressures of from 10.3 bar to 138 bar, and a productcomposed of alcohols and diols is isolated. In the examples, a series ofcatalysts containing copper, zinc and chromium is described.

Finally, the documents cited hereinbelow disclose the use ofchromium-free copper catalysts for hydrogenating GBL to BDO.

A catalyst consisting of CuO and ZnO is described in WO 82/03854. In thegas phase at a pressure of 28.5 bar and a temperature of 217° C., thisachieves a BDO selectivity of 98.4%. However, the conversion of pure GBLis unsatisfactorily low.

Deposited copper catalysts doped with palladium and potassium aredescribed in U.S. Pat. Nos. 4,797,382; 4,885,411 and EP-A 0 318 129.They are suitable for converting GBL to butanediol.

The use of a mixture of GBL and water as the feed stream in combinationwith a copper oxide-zinc oxide catalyst is described in U.S. Pat. No.5,030,773 A. This discloses that the activity of such catalystsincreases when from 1 to 6% of water is admixed into the pure GBL streamand this mixture is hydrogenated in the gas phase. When pure GBL is usedin this reaction, extra water has to be admixed and then has to beremoved again. If GBL were to be used which resulted from hydrogenationof MA, 17% of water would be present in the feed. Accordingly, at least11% of water would have to be removed before hydrogenation to BDO.

JP-A 0 634 567 describes a catalyst comprising copper, iron and aluminumwhich is suitable for hydrogenating pure GBL to BDO at high pressure(250 bar).

A process for preparing BDO starting from maleic esters is cited in WO99/35113. Hydrogenation is effected in three successive stages. Startingfrom maleic esters, succinic ester is prepared over a noble metalcatalyst and is then converted in a second stage to GBL and THF. GBL isremoved and converted to BDO in a third stage at elevated pressure.

WO 99/35114 describes a process for preparing BDO by liquid phasehydrogenation of GBL, succinic esters or mixtures of the two atpressures of from 60 bar to 100 bar and temperatures of from 180° C. to250° C. The catalyst used is a copper oxide-zinc oxide catalyst.

A further gas phase variant of hydrogenation of GBL to BDO is disclosedby WO 99/52845 and uses a copper oxide-zinc oxide catalyst. In additionto the customary reaction feed, carbon monoxide is admixed with thehydrogen in order to coproduce methanol.

EP-A 0 382 050 concerns the hydrogenation of pure GBL over a catalystcomprising cobalt oxide, copper oxide, manganese oxide and molybdenumoxide.

Direct preparation of BDO starting from MA is also known. The documentscited hereinbelow describe this reaction using chromium catalysts.

DE 2 845 905 describes a continuous process for preparing butanediolstarting from maleic anhydride. MA dissolved in monohydric aliphaticalcohols is reacted with hydrogen at pressures of 250 bar and 350 barover copper chromite catalysts.

A process for coproducing BDO and THF starting from MA over copper-,chromium- and manganese-containing catalysts is disclosed by EP-A 0 373947. Mixtures of MA and GBL, mixtures of MA and 1,4-dioxane and pure MAare used. In all cases, mixtures of THF and BDO are obtained. Adisadvantage of this process is the high yields of tetrahydrofuran.

The documents CN-A 1 113 831-A, CN-A 1 116 615-A, CN-A 1 138 018-A andCN-A 1 047 328 disclose chromium catalysts. CN 1 137 944 A uses acopper, chromium, manganese, barium and titanium catalyst.

According to the disclosure of CN-A 1 182 639, a copper, chromium, zincand titanium catalyst can be utilized for hydrogenating mixtures of GBLand MA.

CN-A 1 182 732 describes a process for preparing BDO by gas phasehydrogenation of MA over copper and chromium catalysts at from 200 to250° C. and a pressure of from 30 to 70 bar. MA is hydrogenateddissolved in a suitable solvent.

The documents cited hereinbelow disclose finally the directhydrogenation of MA to BDO using chromium-free catalysts.

For instance, DE-A 2 455 617 describes a three-stage process forpreparing BDO. In a first stage, solutions of MA in GBL are hydrogenatedto SA in GBL over a nickel catalyst. In a second stage at high pressure(80-200 bar) and at relatively high temperature, this solution of SA andGBL is hydrogenated to GBL in the liquid phase, then water, succinicanhydride and succinic acid are removed from GBL and the pure GBL ispartially recycled and converted to butanediol in a third process stageover a copper-zinc oxide catalyst in the liquid phase at high pressure.

In U.S. Pat. No. 4,301,077, a ruthenium catalyst is used to hydrogenateMA to BDO.

DE-A 3 726 510 discloses the use of a catalyst comprising copper, cobaltand phosphorus for direct hydrogenation of MA.

In J0 2025-434-A, a pure copper oxide-zinc oxide catalyst is used.According to the examples, pure MA may be converted at a pressure of 40bar. However, the yield of butanediol is only 53.3 mol %, and asecondary yield of 40.2 mol % of GBL is found.

EP-A 373 946 discloses a process by which gas phase MA is converteddirectly to BDO using a rhenium-doped copper oxide-zinc oxide catalyst.

The coproduction of BDO and TBF is provided by patent applications J02233-627-A (using a copper-zinc-aluminum catalyst), J0 2233-630-A (usinga copper-chromium catalyst comprising manganese, barium and silicon),and J0 2233-631-A (using a catalyst comprising copper and aluminum). Theuse of these catalysts results in the production of large quantities ofTHF as well as BDO from MA as well as BDO.

A catalyst comprising copper, manganese and potassium is described inJ0-A 2233-632.

EP-A 431 923 describes a two-stage process for preparing BDO and THF bypreparing GBL in a first stage by liquid phase hydrogenation of MA andconverting it in a second stage to butanediol by gas phase reaction overa catalyst comprising copper and silicon.

U.S. Pat. No. 5,196,602 discloses a process for preparing butanediol byhydrogenating MA or maleic acid using hydrogen in a two-stage process.In a first stage, MA is hydrogenated to SA and/or GBL which is thenconverted to BDO in a second stage in the presence of an Ru-containingcatalyst.

The technologies on which the above-cited documents are based utilizeprepurified MA, which, after its preparation, has generally been freedof impurities by distillation, as a reactant in the hydrogenationreactions MA is prepared by partial oxidation of certain hydrocarbonsincluding benzene, butene mixtures and also n-butane, and preference isgiven to using the latter. The crude product of the oxidation, inaddition to the desired MA, comprises in particular by-products such aswater, carbon monoxide, carbon dioxide, unconverted starting hydrocarbonand also acetic and acrylic acid, and these by-products are independentof the hydrocarbons used in the oxidation. Normally, the by-products areremoved by complicated processes, for example by distillation, asmentioned above. The purification is necessary in particular because thecatalysts used in the hydrogenation processes are generally sensitive tosuch impurities. The deactivation of the catalysts is a problem evenwhen purified MA is used, since fouling by polymerization productsthereof means that the catalyst generally has to be regenerated atrelatively short intervals which are often about 100 hours. The tendencyto deactivation is increased further when polymerizable impurities, forexample acrylic acid, are present. This fact is known to those skilledin the art and is also described, for example, in patent applicationsEP-A 322 140, WO 91/16132 and DE-A 240 44 93.

Hitherto, only a single document in the prior art discloses thehydrogenation of only coarsely pre-purified MA. WO 97/43234 disclosesthe absorption of maleic anhydride from maleic anhydride-containing gasstreams which stem from the oxidation of hydrocarbons using absorbentswhich boil at least 30° C. higher, stripping the maleic anhydride fromthese absorbents with the aid of hydrogen and hydrogenating the maleicanhydride-containing hydrogen stream in the gas phase over aheterogeneous catalyst. This gives mainly BDO, as well as smallquantities of GBL and THF. The hydrogenation is carried out at fromabout 150° C. to 300° C. and a pressure of from 5 bar to 100 bar in thegas phase. The catalysts used are promoted copper catalysts as describedin Journal of Catalysis 150, pages 177 to 185 (1994). These are chromiumcatalysts of the Cu/Mn/Ba/Cr and Cu/Zn/Mg/Cr type. Accordingly, thisapplication discloses the use of chromium catalysts for hydrogenatinggrades of MA which have the abovementioned impurities. However, the useof chromium catalysts is today avoided as far as possible owing to theirtoxicity.

Owing to their toxicity, novel technologies are moving more and moreaway from the use of chromium catalysts. Examples of chromium-freecatalyst systems can be found in the documents WO 99/35139 (Cu—Znoxide), WO 95/22539 (Cu—Zn—Zr) and also U.S. Pat. No. 5,122,495(Cu—Zn—Al oxide).

In the field of MA hydrogenation to subsequent products, in particularGBL, THF and/or BDO, there is thus a virtually limitless prior art, andonly a selection of the entire existing prior art has been cited above.

To summarize, it can be said that the technical problems occurring inpreparing BDO by hydrogenating MA have been solved in that satisfactoryyields and selectivities for BDO, i.e. only insignificant formation ofTHF, have been achieved. This has been achieved by different measures orthe combination of different measures.

In general, BDO has been obtained by direct hydrogenation of pure GBLwhich has been obtained in turn by the hydrogenation of MA andsubsequent costly and inconvenient purification. In every case, thereactant used was pure MA which only contains small quantities ofimpurities, since otherwise no satisfactory selectivity and catalyston-stream time could be achieved. Chromium catalysts were used, inparticular in the second stage, in order to achieve high BDO selectivityand the desired on-stream time. To avoid the use of chromium catalysts,there is the alternative of using noble metal catalysts which, in termsof yield, selectivity and also durability, are comparable with chromiumcatalysts, but are distinctly more costly.

The preferred procedure of the reaction in two separate stages alsoinvolves costly and inconvenient purification of the GBL after the firsthydrogenation stage in order to achieve a long on-stream time of thecatalysts, in particular with regard to the desired selectivity.Hitherto, the abovementioned WO 97/43234 discloses the only processwhich uses only coarsely prepurified MA as the reactant for preparingBDO by hydrogenation. The process is carried out in one stage, and soavoids the workup after the first hydrogenation stage. However, onlychromium catalysts are suitable for this conversion.

It is an object of the present invention to provide a process forpreparing BDO from MA which does not use pure MA and requires at leastno costly and inconvenient purification of the first stage reactionproducts and also provides very good butanediol selectivities andcatalyst on-stream times. The process shall further require no chromiumcatalysts, and preferably also no catalysts which comprise noble metals,and shall also have a high selectivity for BDO, and in particulardeliver little THF.

We have found that this object is achieved by a process for preparingoptionally alkyl-substituted 1,4-butanediol by two-stage catalytichydrogenation in the gas phase of C₄-dicarboxylic acids and/or ofderivatives thereof having the following steps:

-   -   a) introducing a gas stream of a C₄-dicarboxylic acid or of a        derivative thereof at from 200 to 300° C. and from 2 to 60 bar        into a first reactor and catalytically hydrogenating it in the        gas phase to a product which contains mainly optionally        alkyl-substituted γ-butyrolactone;    -   b) removing succinic anhydride from the product stream obtained        in step a);    -   c) introducing the product stream obtained in step b) into a        second reactor at a temperature of from 150° C. to 240° C. and a        pressure of from 15 to 100 bar and catalytically hydrogenating        it in the gas phase to optionally alkyl-substituted        1,4-butanediol;    -   d) removing the desired product from intermediates, by-products        and any unconverted reactant;    -   e) optionally recycling unconverted intermediates into one or        both hydrogenation stages,        -   said hydrogenation stages each using a catalyst which            comprises ≦95% by weight, preferably from 5 to 95% by            weight, in particular from 10 to 80% by weight, of CuO, and            ≧5% by weight, preferably from 5 to 95% by weight, in            particular from 20 to 90% by weight, of an oxidic support,            and said second reactor having a higher pressure than said            first reactor.

To achieve the desired BDO selectivities, the maintenance of certainreaction parameters in both hydrogenation stages is necessary, and theseparameters are cited hereinbelow.

It has been found that SA contributes substantially to fast deactivationof the hydrogenation catalyst used in step c). The process according tothe invention which contemplates the removal of SA accordingly provideshigh BDO yields accompanied by long catalyst on-stream times and verygood BDO selectivities.

The process according to the invention contemplates no costly andinconvenient purification of the product stream leaving the firsthydrogenation stage a), since SA, in particular owing to its highboiling point, can be removed from the product stream withoutcomplicated apparatus. Comparatively simple, inexpensive and convenientmeasures thus allow the selectivity and on-stream time of the catalystto be improved. Preference is given to removing the SA by simplecondensing out in a partial condensation. The removal process isdescribed in greater detail below.

In the process according to the invention, reactants of differing puritymay be used in the hydrogenation reaction. It will be appreciated that areactant of high purity, in particular MA, may be used in thehydrogenation reaction. However, an advantage of the process accordingto the invention is further that the use of reactants, in particular MA,which is contaminated with the customary compounds resulting from theoxidation, i.e. benzene, butenes or n-butane, and also any furthercomponents, is also made possible. Accordingly, the hydrogenationprocess according to the invention in a further embodiment may comprisea preceding stage which comprises the preparation of the reactant to behydrogenated by partial oxidation of a suitable hydrocarbon and also theremoval of the reactant to be hydrogenated from the product streamobtained. Preference is given to carrying out only a coarse removalwhich requires no unnecessary effort and allows a quantity of impuritiesto remain in the reactant which was intolerable by the prior artprocesses.

In particular, this reactant to be hydrogenated is MA. Preference isgiven to using MA which stems from the partial oxidation ofhydrocarbons. Useful hydrocarbons include benzene, C₄-olefins (forexample n-butene, C₄-raffinate streams) or n-butane. Particularpreference is given to using n-butane, since it is an inexpensive,economical starting material. Processes for the partial oxidation ofn-butane are described, for example, in Ullmann's Encyclopedia ofIndustrial Chemistry, 6^(th) Edition, Electronic Release, maleic andfumaric acids—maleic anhydride.

Preference is given to then taking up the reaction effluent obtained inthis manner in a suitable organic solvent or solvent mixture which has aboiling point at atmospheric pressure which is at least 30° C. higherthan that of MA.

This solvent (absorbent) is brought to a temperature in the range from20 to 160° C., preferably from 30 to 80° C. The maleicanhydride-containing gas stream from the partial oxidation may becontacted with the solvent in various ways: (i) passing the gas streaminto the solvent (for example via gas inlet nozzles or sparging rings),(ii) spraying the solvent into the gas stream and (iii) countercurrentcontact between the gas stream flowing upward and the solvent flowingdownward in a tray or packed column. In all three variants, apparatusknown to those skilled in the art may be used for gas absorption. Whenchoosing the solvent to be used, care must be taken that it does notreact with the reactant, for example the preferably used MA. Usefulsolvents are: tricresyl phosphate, dibutyl maleate, butyl maleate, highmolecular weight waxes, aromatic hydrocarbons having a molecular weightof from 150 to 400 and a boiling point above 140° C., for exampledibenzylbenzene, alkyl phthalates and dialkyl phthalates havingC₁-C₁₈-alkyl groups, for example dimethyl phthalate, diethyl phthalate,dibutyl phthalate, di-n-propyl and diisopropyl phthalate, undecylphthalate, diundecyl phthalate, methyl phthalate, ethyl phthalate, butylphthalate, n-propyl or isopropyl phthalate; di-C₁-C₄-alkyl esters ofother aromatic and aliphatic dicarboxylic acids, for exampledimethyl-2,3-naphthalenedicarboxylate, dimethyl-1,4-cyclohexanedicarboxylate; C₁-C₄-alkyl esters of other aromatic and aliphaticdicarboxylic acids, for example methyl-2,3-naphthalenedicarboxylates,methyl-1,4-cyclohexanedicarboxylate; methyl esters of long-chain fattyacids having for example from 14 to 30 carbon atoms, high-boilingethers, for example dimethyl ethers of polyethylene glycol, for exampletetraethylene glycol dimethyl ether.

Preference is given to the use of phthalates.

The solution resulting from treatment with the absorbent generally hasan MA content of from about 5 to 400 grams per liter.

The gas stream remaining after treatment with the absorbent containsmainly the by-products of the preceding partial oxidation such as water,carbon monoxide, carbon dioxide, unconverted butanes, acetic acid andacrylic acid. The offgas stream is virtually free of MA.

The dissolved MA is then stripped from the absorbent. This is effectedusing hydrogen at or at a maximum of 10% above the pressure of thesubsequent hydrogenation or alternatively under reduced pressure withsubsequent condensation of remaining MA. In the stripping column, atemperature profile is observed which results from the boiling points ofMA at the top and the virtually MA-free absorbent at the bottom of thecolumn at the column pressure in each case and the dilution with carriergas (in the first case with hydrogen) used.

In order to avoid solvent losses, rectifying internals may be disposedabove the feed of the crude MA stream. The virtually MA-free absorbenttaken off at the bottom is fed back into the absorption zone. The H₂/MAratio is from about 20 to 400. Otherwise, the condensed MA is pumpedinto an evaporator and evaporated there into the cycle gas stream.

The MA-hydrogen stream also contains by-products which result from thepartial oxidation of n-butane, butenes or benzene usingoxygen-containing gases, and also unremoved absorbent. The by-productsare in particular acetic acid and acrylic acid, and there is also water,maleic acid and also dialkyl phthalates which are preferably used asabsorbents. The MA contains acetic acid in quantities from 0.01 to 1% byweight, preferably from 0.1 to 0.8% by weight, and acrylic acid inquantities from 0.01 to 1% by weight, preferably from 0.1 to 0.8% byweight, based on MA. In the hydrogenation stage, acetic acid and acrylicacid are partially or completely hydrogenated to ethanol and propanolrespectively. The maleic acid content is from 0.01 to 1% by weight, inparticular from 0.05 to 0.3% by weight, based on MA.

When alkyl and dialkyl phthalates are used as absorbents, the contentthereof in the MA depends strongly on the correct operation of thestripping column, in particular of the rectifying section. Phthalatecontents of up to 1.0% by weight, in particular of up to 0.5% by weight,should not be exceeded when the column is operated in a suitable manner,since the consumption of absorbents otherwise becomes too high.

The hydrogen/maleic anhydride stream which is preferably obtained asdescribed above is then fed into the first hydrogenation zone andhydrogenated. The catalyst activities and on-stream times are virtuallyunchanged compared to the use of extensively, for example bydistillation, prepurified MA.

The gas stream leaving the first reactor is freed of SA and may then,according to the invention, be further processed in various ways. The SAmay be removed by measures known to those skilled in the art, forexample by partial condensation, optionally in countercurrent,condensation or distillation. The above-described measures distinctlyreduce the SA content of the gas stream. The acceptable residual SAcontent for the process varies and depends on many different factors,for example the composition of the catalyst in the second hydrogenationstage. It is frequently desirable to attain a residual SA content of <about 0.3 to 0.2% by weight. This value is achieved in particular whenthe process in the first hydrogenation stage is carried out in such amanner that the SA content of the exit stream is about 1% by weight.

According to one variant, the gas stream freed of SA is compressed tothe higher pressure of the second hydrogenation stage and fed in thisform with any recycled GBL to the second hydrogenation.

In a further variant, the gas stream may be cooled to from 10 to 60° C.The reaction products are condensed out and passed into a separator. Theuncondensed gas stream may be removed and, preferably after feeding intoa cycle gas compressor, returned to the first hydrogenation cycle.By-products formed in the recycled cycle gas stream may be removed bymeasures to those skilled in the art, preferably by bleeding off a smallquantity of cycle gas. The reaction products which have condensed outare withdrawn from the system and introduced into the secondhydrogenation cycle. The reaction products are brought into the gasphase there under pressure with any recycled GBL and contacted with thesecond catalyst. Any recycled GBL may also be directly introduced intothe second hydrogenation reactor in gaseous form.

In a further variant, the GBL-laden gas stream from the first stage iscompressed to the pressure of the second stage and the recycled gas ofthe second stage is expanded into the entrance of the first stage,optionally while performing work.

In all reaction variants, the gas stream leaving the second reactor iscooled, preferably to from 10 to 60° C. The reaction products arecondensed out and passed into a separator. The uncondensed gas stream istaken from the separator and fed into the cycle gas compressor. A smallcycle gas quantity is bled off. Preference is given to continuouslywithdrawing the reaction products which have condensed out from thesystem and feeding them to a workup. The by-products in the liquid phasewhich has condensed out are mainly THF and n-butanol, as well as smallquantities of propanol.

The by-products and also water and the desired product BDO are thenisolated from the liquid hydrogenation residue of the second stage. Thisis generally effected by fractional distillation. By-products andintermediates, for example GBL and di-BDO, may be returned to thehydrogenation of the first and/or second stage, preferably of the secondstage, or alternatively worked up distillatively.

The process according to the invention may be carried out batchwise,semicontinuously or continuously. Preference is given to carrying it outcontinuously.

An important parameter is the maintenance of a suitable reactiontemperature in both hydrogenation stages.

In the first hydrogenation stage, preference is given to obtaining thisby a sufficiently high temperature of the reactants when entering thefirst hydrogenation reactor. This starting hydrogenation temperature isfrom 200 to 300° C., preferably from 235 to 270° C. In order to obtainthe desired selectivity and yield in the first stage, the reactionshould preferably be carried out in such a manner that the catalyst bedwhere the actual reaction takes place is at a suitably high reactiontemperature. This hot spot temperature is set after the reactants enterthe reactor and is preferably from 210 to 310° C., in particular from245 to 280° C. Preference is given to carrying out the process in such amanner that the entrance temperature and the exit temperature of thereaction gases are below this hot spot temperature. The hot spottemperature is advantageously in the first half of the reactor, inparticular when it is a tube bundle reactor. The hot spot temperature ispreferably from 5 to 30° C., in particular from 5 to 15° C., morepreferably from 5 to 10C., above the entrance temperature. When thehydrogenation is carried out below the minimum entrance and hot spottemperatures and MA is used as the reactant, the quantity of SAgenerally increases while at the same time the GBL and BDO quantitiesdecrease. Such a temperature also results in the observation of catalystdeactivation in the course of the hydrogenation due to fouling bysuccinic acid, fumaric acid and/or SA and mechanical damage to thecatalyst. In contrast, when MA is used as the reactant above the maximumentrance and hot spot temperatures, the BDO yield and selectivitygenerally fall to unsatisfactory values. Increased formation of THF,n-butanol and n-butane is observed, i.e. the products of furtherhydrogenation.

In the second hydrogenation stage, the entrance temperature (startinghydrogenation temperature) is from 150° C. to 260° C., preferably from175° C. to 225° C., in particular from 180 to 200° C. When thehydrogenation is carried out below the minimum entrance temperature, thequantity of BDO formed falls. The catalyst loses activity. Below theminimum temperature, condensation of the starting materials and damageto the copper catalyst by water are also to be expected. In contrast,when GBL is used as the reactant for hydrogenation above the maximumentrance temperature, the BDO yield and selectivity fall tounsatisfactory values. At these temperatures, the hydrogenationequilibrium between BDO and GBL is on the side of GBL so that lessconversion is obtained, but increased by-product formation byoverhydrogenation to THF, n-butanol and n-butane is observed atrelatively high temperatures.

The temperature increase of the gas stream in the reactor should notexceed 110° C., preferably 40° C., and in particular should not be morethan 20° C. Large temperature increases here also lead tooverhydrogenation reactions and (BDO+GBL) selectivity loss.

In the first hydrogenation stage, a pressure of from 2 to 60 bar,preferably a pressure of from 2 to 20 bar and more preferably a pressureof from 5 to 15 bar, is selected. In this pressure range, thehydrogenation of MA proceeds with very substantial suppression of THFformation from the initially formed intermediate GBL.

In the second hydrogenation stage, a pressure of from 15 to 100 bar,preferably a pressure of from 35 to 80 bar and more preferably apressure of from 50 to 70 bar, is chosen. At the temperature specifiedfor the second hydrogenation stage, the conversion of GBL to BDOincreases with pressure. Higher pressure accordingly enables a lower GBLrecycling rate to be chosen. The pressure in the second hydrogenationstage is above the pressure of the first hydrogenation stage.

The catalyst space velocity of the first hydrogenation stage ispreferably in the range from 0.02 to 1, in particular from 0.05 to 0.5,kg of reactant/l of catalyst □ hour. In the case of MA, when thecatalyst hourly space velocity of the first stage increases above thisrange, an increase of the SA and succinic acid contents in thehydrogenation effluent is observed. The catalyst hourly space velocityof the second hydrogenation stage is in the range from 0.02 to 1.5, inparticular from 0.1 to 1, kg of reactant/l of catalyst □ hour. When thecatalyst hourly space velocity is increased above this range, incompleteconversion of GBL is to be expected. This may optionally be compensatedfor by an increased recycling rate, although it will be appreciated thatthis is not preferred.

The hydrogen/reactant molar ratio is also a parameter which influencesthe product distribution and also the economic viability of the processaccording to the invention. From an economic point of view, a lowhydrogen/reactant ratio is desirable. The lower limit is at a value of5, although higher hydrogen/reactant molar ratios of from 20 to 600 aregenerally used. The use of catalysts used according to the invention andalso the maintenance of the temperatures specified according to theinvention allows the use of advantageous, low hydrogen/reactant ratiosin the first stage hydrogenation which are preferably from 20 to 200,preferably from 40 to 150. The most advantageous range is from 50 to100.

In order to attain the hydrogen/reactant molar ratios used according tothe invention, a portion, advantageously the majority, of the hydrogenis customarily recycled in both the first and the second hydrogenationstages. To this end, the cycle gas compressor familiar to those skilledin the art is generally used. The hydrogen quantity consumed byhydrogenation is replaced. In a preferred embodiment, a portion of thecycle gas is bled off, in order to remove inert compounds, for examplen-butane. The recycled hydrogen may also, optionally after preheating,be utilized to evaporate the reactant stream.

Together with the hydrogen cycle gas, all products are recycled which donot condense out or do so incompletely when the gas streams leaving thehydrogenation reaction are cooled. These are in particular THF, waterand by-products such as methane and butane. The cooling temperature ispreferably from 0 to 60° C., preferably from 20 to 45° C.

Useful reactor types include all apparatus suitable for heterogeneouslycatalyzed reactions having gaseous reaction and product streams.Preference is given to tubular reactors, shaft reactors or reactorshaving internal heat removal means, for example tube bundle reactors,and the use of a fluidized bed is also possible. Particular preferenceis given to using tube bundle reactors for the first hydrogenationstage, and to shaft reactors for the second hydrogenation stage. In boththe first and the second hydrogenation stage, more than one reactorconnected in parallel or in series may be used. In principle, there mayalso be intermediate feeding between the catalyst beds. It is alsopossible to provide intermediate cooling between or in the catalystbeds. When fixed bed reactors are used, dilution of the catalyst byinert material is possible.

An important point of the present invention is the choice of thecatalysts for both stages which have copper oxide as the catalyticallyactive main component. This is applied to an oxidic support which mayonly have a small number of acidic sites. When a catalyst having toohigh a number of acidic sites is used, BDO is dehydrated and THF isformed.

A suitable support material which has a sufficiently low number ofacidic sites is a material selected from the group of ZnO, Al₂O₃, SiO₂,TiO₂, ZrO₂, CeO₂, MgO, CaO, SrO, BaO and Mn₂O₃ and mixtures thereof.Preferred support materials are ZnO/Al₂O₃ mixtures, the delta-, theta-,alpha- and eta-modifications of Al₂O₃ and also mixtures which compriseat least one component each firstly from the group of SiO₂, TiO₂, ZrO₂,and secondly from the group of ZnO, MgO, CaO, SrO and BaO. Particularlypreferred support materials are pure ZnO, ZnO/Al₂O₃ mixtures in a weightratio of from 100:1 to 1:2 and mixtures of SiO₂ with MgO, CaO and/or ZnOin a weight ratio of 200:1 to 1:1.

The copper oxide quantity is ≦95% by weight, preferably from 5 to 95% byweight, in particular from 15 to 80% by weight; the support is used inquantities of ≧5% by weight, preferably from 5 to 95% by weight, inparticular from 20 to 85% by weight.

Owing to the toxicity of chromium catalysts, preference is giving tousing chromium-free catalysts. It will be appreciated that correspondingchromium catalysts known to those skilled in the art are technicallyalso suitable for use in the process according to the invention,although the desired advantages which are in particular of environmentaland technical nature do not accrue.

The same catalyst may be used in both hydrogenation stages, butpreference is given to the use of different catalysts.

Optionally, the catalysts used according to the invention may compriseone or more further metals or a compound thereof, preferably an oxide,from groups 1 to 14 (IA to VIIIA and IB to IVB of IUPAC nomenclature) ofthe Periodic Table. When a further metal is used, preference is given tousing Pd in quantities of ≦1% by weight, preferably ≦0.5% by weight, inparticular ≦0.2% by weight. However, the use of a further metal or metaloxide is not preferred.

The catalysts used may additionally contain an auxiliary in a quantityof from 0 to 10% by weight. Auxiliaries are organic and inorganicmaterials which contribute to improved processing during catalystpreparation and/or to an increase in the mechanical stability of thecatalyst shaped bodies. Useful auxiliaries are known to those skilled inthe art; examples include graphite, stearic acid, silica gel and copperpowder.

The catalysts can be prepared by methods known to those skilled in theart. Preference is given to processes which provide the copper oxidefinely divided and intimately mixed with the other components, greaterpreference to impregnation and precipitation reactions.

These starting materials may be processed by known methods to give theshaped bodies, for example extrusion, tableting or agglomerationprocesses, optionally with the use of auxiliaries.

Alternatively, catalysts according to the invention may be prepared, forexample, by applying the active component to a support, for example bycoating or vapor deposition. Catalysts according to the invention mayalso be obtained by shaping a heterogeneous mixture of active componentor precursor compound thereof with a support component or precursorcompound thereof.

The hydrogenation according to the invention which may, as well as MA,use other, above-defined C₄-dicarboxylic acids or derivatives thereof asreactants uses the catalyst in reduced, activated form. The catalyst isactivated using reducing gases, preferably hydrogen or hydrogen/inertgas mixtures either before or after installation in the reactors wherethe process according to the invention is carried out. If the catalysthas been installed in oxidic form in the reactor, it may be activatedeither before startup of the plant with the hydrogenation according tothe invention or else during the startup, i.e. in situ. Separateactivation before plant startup is generally effected using reducinggases, preferably hydrogen or hydrogen/inert gas mixtures at elevatedtemperatures, preferably from 100 to 350° C. In situ activation iseffected when starting up the plant by contacting with hydrogen atelevated temperature.

Preference is given to using the catalysts in the form of shaped bodies.Examples include extrudates, ribbed extrudates, other extrudate forms,tablets, rings, spheres and spall.

The BET surface area of the copper catalysts in the oxidic state shouldbe from 10 to 300 m²/g, preferably from 15 to 175 m²/g, in particularfrom 20 to 150 m²/g. The copper surface area (N₂O decomposition) of thereduced catalyst in the installed state should be >0.2 m²/g,preferably >1 m²/g, in particular >2 m²/g.

In one variant of the invention, catalysts are used which have a definedporosity. The shaped bodies of these catalysts have a pore volume of≧0.01 ml/g for pore diameters of >50 nm, preferably of ≧0.025 ml/g forpore diameters of >100 nm and in particular of ≧0.05 ml/g for porediameters of >200 nm. The ratio of macropores having a diameter of >50nm to the total pore volume for pores having a diameter of >4 nm isalso >10%, preferably >20%, in particular >30%. The porosities mentionedwere determined by mercury intrusion according to DIN 66133. The datawere evaluated in the pore diameter range of from 4 nm to 300 μm.

The catalysts used according to the invention generally have sufficienton-stream time. However, in the event that the activity and/orselectivity of the catalyst should fall in the course of its operatingtime, it may be regenerated by methods known to those skilled in theart. These include preferably reductive treatment of the catalyst in ahydrogen stream at elevated temperature. The reductive treatment mayoptionally be followed by oxidative treatment. To this end, a molecularoxygen-containing gas mixture, for example air, is passed through thecatalyst bed at elevated temperature. There is also the possibility ofwashing the catalysts with a suitable solvent, for example ethanol, THF,BDO or GBL, and then drying them in a gas stream.

The process according to the invention is illustrated by the exampleshereinbelow.

EXAMPLE 1ss

a) Catalyst Preparation

b) Catalyst Activation

Before the beginning of the reaction, the catalyst is subjected to ahydrogen treatment in the hydrogenation apparatus. The reactor is heatedto 180° C. and the catalyst activated for the time specified in Table 1using the mixture of hydrogen and nitrogen specified in each case atatmospheric pressure.

TABLE 1 Time Hydrogen Nitrogen (minutes) (l/h, stp) (l/h, stp) 120 10550 30 25 400 15 60 100 180 60 0

c) Hydrogenation Apparatus

The pressure apparatus used for the hydrogenation reaction consists ofan evaporator, a reactor, a condenser having a quench feed, a hydrogenfeed, an offgas line and a cycle gas blower. The pressure in theapparatus is held constant.

The melted MA is pumped from above onto the preheated (245° C.)evaporator and evaporated. A mixture of fresh hydrogen and cycle gaslikewise reaches the evaporator from above. In this manner, hydrogen andMA pass into the heated reactor from below. The reactor contents consistof a mixture of glass rings and catalyst. After the hydrogenation, theTHF formed together with water, other reaction products and hydrogenleaves the reactor and is precipitated in the condenser by quenching. Aportion of the cycle gas is bled off before the remainder, mixed withfresh hydrogen, reenters the evaporator.

The condensed liquid reaction effluent, the offgas and the cycle gas arequantitatively analyzed by gas chromatography.

EXAMPLE 1d

d) Hydrogenation of Maleic Anhydride Obtained from N-Butane

The reactor of the hydrogenation apparatus described in Example 1b ischarged with 220 ml of the catalyst prepared according to Example 1a and130 ml of glass rings. The activation was effected as described inExample 1b.

The reactant used is maleic anhydride obtained from n-butane whichcontains 500 ppm of acrylic acid, 1500 ppm of acetic acid and 100 ppm ofdibutyl phthalate. The reaction is carried out for 1000 h. Over theentire period, no deactivation of the catalyst, i.e. no reduction in themaleic anhydride conversion and/or tetrahydrofuran yield, is observed.Butanediol is not observed by gas chromatography. The hydrogenationreaction parameters and the results are summarized in Table 2.

EXAMPLE 1

a) Catalyst Preparation

From a metal salt solution containing copper nitrate and zinc nitrate,sodium carbonate is used at 50° C. and a pH of around 6.2 to precipitatea mixed basic metal carbonate. The metal salt solution used contains themetals corresponding to a catalyst composition of 70% of CuO and 30% ofZnO.

The precipitate is filtered, washed, dried, calcined at 300° C. andpressed with 3% by weight of graphite to give tablets each of height anddiameter 3 mm.

b) Catalyst Activation

Before the beginning of the reaction, the catalyst is subjected to ahydrogen treatment in the hydrogenation apparatus. The reactor is heatedto 180° C. and the catalyst activated for the time specified in Table 1using the mixture of hydrogen and nitrogen specified in each case atatmospheric pressure.

TABLE 1 Time Hydrogen Nitrogen (Minutes) (l/h, stp) (l/h, stp) 120 10550 30 25 400 15 60 100 180 60 0

c) Hydrogenation Apparatus

The pressure apparatus used for the hydrogenation reaction consists ofan evaporator, a reactor, a condenser having a quench feed, a hydrogenfeed, an offgas line and a cycle gas blower. The pressure in theapparatus is held constant.

The melted MA is pumped from above onto the preheated (245° C.)evaporator and evaporated. A mixture of fresh hydrogen and cycle gaslikewise reaches the evaporator from above. In this manner, hydrogen andMA pass into the heated reactor from below. The reactor contents consistof a mixture of glass rings and catalyst. After the hydrogenation, theGBL formed together with water, other reaction products and hydrogenleaves the reactor and is precipitated in the condenser by quenching. Aportion of the cycle gas is bled off before the remainder, mixed withfresh hydrogen, reenters the evaporator.

The condensed liquid reaction effluent, the offgas and the cycle gas arequantitatively analyzed by gas chromatography.

At a reactor temperature of 255° C. a pressure of 5 bar and a catalysthourly space velocity of 0.27 kg/L_(cat)h at a hydrogen: MA molar ratioof 85:1, a reaction effluent of the composition: 91% of GBL, 5% of THF,1% of BDO and 1% of SA

EXAMPLE 2

a) Catalyst Preparation

From an aqueous solution of zinc nitrate and aluminum nitrate, sodiumcarbonate solution is used at 50° C. and a pH of 6.8 to precipitate asolid having the composition 64% of ZnO and 36% of Al₂O₃ (based on 100%oxide), filtered and washed. The filtercake was dried and calcined at425° C. for one hour.

The above-described support is added to a nitric acid solution of coppernitrate and zinc nitrate (metal ratio corresponding to 16.6% by weightof CuO and 83.4% of ZnO) and mixed intensively at 70° C. From thismixture, sodium carbonate solution is used to precipitate a solid at 70°C. and a pH of 7.4 and the suspension is stirred at constant temperatureand pH for a further 2 h. The solid is filtered off, washed, dried andcalcined at 430° C. for one hour. The catalyst powder obtained in thismanner was mixed with 1.5% by weight of graphite and 5% by weight ofcopper powder and pressed to give tablets of diameter 1.5 mm and height1.5 mm. The tablets were finally calcined at 330° C. for 1 h, and had aside crushing strength of 50N and a chemical composition of 66% ofCuO/24% of ZnO/5% of Al₂O₃/5% of Cu.

b) Catalyst Activation Similar to Example 1b

c) Hydrogenation Apparatus

The reactor of the hydrogenation apparatus described in Example 1c ischarged with 220 ml of the catalyst prepared according to Example 2a and130 ml of glass rings. The activation was effected as described inExample 1b.

The reactant used is the reaction effluent of the MA hydrogenation ofExample 1 from which more than 50% of the SA content have been removedby partial condensation. At a reactor temperature of 180° C., a pressureof 60 bar and a catalyst hourly space velocity of 0.15 kg/L_(cat)h(hydrogen: GBL molar ratio 200:1), a reaction effluent is obtainedhaving the composition: 87% of BDO, 7% of GBL, 5% of THF.

1. A process for preparing optionally alkyl-substituted 1,4-butanediolby two-stage catalytic hydrogenation in the gas phase of C₄-dicarboxylicacids and/or of derivatives thereof having the following steps: a)introducing a gas stream of a C₄-dicarboxylic acid or of a derivativethereof at from 200 to 300° C. and from 2 to 60 bar into a first reactorand catalytically hydrogenating it in the gas phase to a product whichcontains mainly optionally alkyl-substituted γ-butyrolactone; b)removing succinic anhydride from the product obtained in step a); c)introducing the product stream obtained in step b) into a second reactorat a temperature of from 150° C. to 240° C. and a pressure of from 15 to100 bar and catalytically hydrogenating it in the gas phase tooptionally alkyl-substituted 1,4-butanediol; d) removing the desiredproduct from intermediates, by-products and any unconverted reactant; e)optionally recycling unconverted intermediates into one or bothhydrogenation stages, said hydrogenation stages each using a catalystwhich is free from chromium which comprises ≦95% by weight of CuO, and≧5% by weight of an oxidic support, and said second reactor having ahigher pressure than said first reactor, and wherein the catalyst spacevelocity of the first hydrogenation stage is in the range from 0.02 to 1kg of reactant/1 catalyst•hour, and the catalyst space velocity of thesecond hydrogenation stage is in the range from 0.02 to 1.5 kg ofreactant/1 of catalyst•hour.
 2. A process as claimed in claim 1, whereinthe entrance temperature into the first reactor is from 235 to 270° C.and the entrance temperature into the second reactor is from 175° C. to225° C.
 3. A process as claimed in claim 1, wherein the catalytichydrogenation in the first reactor has a hot spot temperature of from210 to 310° C., and the process is carried out in such a manner that thehot spot temperature is above the entrance temperature and the exittemperature of the reaction gases, and is from 5 to 30° C. above theentrance temperature.
 4. A process as claimed in claim 1, wherein thepressure in the first hydrogenation stage is from 2 to 20 bar and thepressure in the second hydrogenation stage is from 35 to 80 bar.
 5. Aprocess as claimed in claim 1, wherein the hydrogen/reactant molar ratioin both reaction stages is >5.
 6. A process as claimed in claim 5,wherein the hydrogen/reactant ratio in the first stage hydrogenation isfrom 20 to
 200. 7. A process as claimed in claim 1, wherein the reactorsused are selected from the group consisting of tubular reactors, shaftreactors, reactors having internal heat removal means, tube bundlereactors and fluidized bed reactors.
 8. A process as claimed in claim 7,wherein the tube bundle reactor is used in the first hydrogenationstage.
 9. A process as claimed in claim 7, wherein a shaft reactor isused in the second hydrogenation stage.
 10. A process as claimed inclaim 1, wherein more than one reactor connected in parallel or inseries is used in the first and/or second hydrogenation stage.
 11. Aprocess as claimed in claim 1, wherein the support material of thecatalyst is selected from the group of ZnO, Al₂O₃, SiO₂, TiO₂, ZrO₂,CeO₂, MgO, GaO, SrO, BaO and Mn₂O₃ and mixtures thereof.
 12. A processas claimed in claim 11, wherein the support material of the catalyst isselected from the group of ZnO/Al₂O₃ mixtures, the delta-, theta-,alpha- and eta- modifications of Al₂O₃ and also mixtures which compriseat least one component each firstly from the group of SiO₂, TiO₂, ZrO₂,and secondly from the group of ZnO, MgO, GaO, SrO and BaO.
 13. A processas claimed in claim 11, wherein the support material is selected fromZnO, ZnO/Al₂O₃ mixtures in a weight ratio of from 100:1 to 1:2 andmixtures of SiO₂ with at least one of MgO, CaO and ZnO in a weight ratioof 200:1 to 1:1.
 14. A process as claimed in claim 1, wherein thecatalyst comprises one or more further metals or a compound of one ormore further metals from groups 1 to 14 of the Periodic Table.
 15. Aprocess as claimed in claim 1, wherein the catalyst is used in the formof shaped bodies.
 16. A process as claimed in claim 1, wherein the BETsurface area of the copper catalysts in the oxidic state is from 10 to300 m²/g.
 17. A process as claimed in claim 1, wherein the coppersurface area of the reduced catalyst in the installed state is >0.2m²/g.
 18. A process as claimed in claim 1, wherein the catalyst used inthe first and second reactors are identical or different.
 19. A processas claimed in claim 1, wherein the shaped bodies of the catalyst usedhave a pore volume of ≧0.01 ml/g for pore diameters of >50 nm.
 20. Aprocess as claimed in claim 1, wherein the ratio of micropores having adiameter of >50 nm to the total pore volume for pores having a diameterof >4 nm is >10%.
 21. A process as claimed in claim 1, wherein thereactant used in the reaction is maleic anhydride.
 22. A process asclaimed in claim 1, wherein maleic anhydride is used which has beenprepared by oxidizing benzene, C₄-olefins or n-butane, and the crudemaleic anhydride obtained by oxidation has been extracted from the crudeproduct mixture using an absorbent and then stripped from this absorbentusing hydrogen.
 23. A process as claimed in claim 22, wherein theabsorbent is selected from the group consisting of tricresyl phosphate,dibutyl maleate, high molecular weight waxes, aromatic hydrocarbonshaving a molecular weight of from 150 to 400 and a boiling point above140° C.
 24. A process as claimed in claim 23, wherein the absorbent isselected from the group consisting of dibenzene, di-C₁-C₄ alkyl estersof aromatic and aliphatic dicarboxylic acids, methyl esters oflong-chain fatty acids having from 14 to 30 carbon atoms, high boilingethers, dimethyl ethers of polyethylene glycol and alkyl phthalates anddialkyl phthalates having C₁-C₁₈ alkyl groups.
 25. A process as claimedin claim 23, wherein the absorbent is selected from the group consistingof dimethyl-2,3-naphthalene dicarboxylate, dimethyl-1,4-cyclohexanedicarboxylate, tetraethylen glycol, dimethyl phthalate, diethylphthalate, dibutyl phthalate, di-n-propyl and di-isopropyl phthalate,undecyl phthalate, diundecyl phthalate, methyl phthalate, ethylphthalate, butyl phthalate, n-propyl phthalate and isopropyl phthalate.26. A process as claimed in claim 22, wherein the maleic anhydride isstripped from the absorbent under reduced pressure or pressures whichcorrespond to the hydrogenation pressure or are a maximum of 10% abovethis pressure.
 27. A process as claimed in claim 1, which is carried outbatchwise, semicontinuously or continuously.
 28. A process as claimed inclaim 1, wherein the SA is removed by partial condensation.
 29. Aprocess as claimed in claim 1, wherein the succinic anhydride is removedfrom the product obtained in step a) to a residual level of from < about0.3 to 0.2% by weight.
 30. A process as claimed in claim 14, wherein thecompound of the one or more further metals is an oxide.